Hydroformylation

ABSTRACT

Olefins having at least 6 carbon atoms are hydroformylated in the presence of a homogeneous catalyst in a continuous process in which a) a vertical tall cylindrical reactor ( 1 ) whose interior space is divided by means of internals ( 2 ) into at least two reaction chambers which extend essentially in the longitudinal direction of the reactor is used, b) at least one olefin is introduced into the reactor together with synthesis gas at the lower end of the first reaction chamber, c) a partially reacted reaction mixture is conveyed from the upper end of a reaction chamber to the lower end of a next reaction chamber; and d) the hydroformylated olefin is taken off at the upper end of the last reaction chamber. The process allows a high conversion at a given reactor volume.

The present invention relates to a process for the hydroformylation ofolefins having at least 6 carbon atoms and to an apparatus for carryingout the process.

Hydroformylation or the oxo process is an important industrial processand is employed for preparing aldehydes from olefins, carbon monoxideand hydrogen. These aldehydes can, if desired, be hydrogenated by meansof hydrogen to produce the corresponding alcohols in the same processstep or subsequently in a separate hydrogenation step. Hydroformylationis carried out in the presence of catalysts which are homogeneouslydissolved in the reaction medium. Catalysts used are generally thecarbonyl complexes of metals of transition group VIII, in particular Co,Rh, Ir, Pd, Pt or Ru, which may be unmodified or modified with, forexample, amine- or phosphine-containing ligands. A summary of processesemployed in industry may be found in J. Falbe, “New Syntheses withCarbon Monoxide”, Springer Verlag 1980, p. 162 ff.

While short-chain olefins can be hydroformylated relatively easily, thehydroformylation of higher olefins sometimes presents problems. This isdue to the decrease in the reaction rate as the chain length increasesand the proportion of internal, nonterminal double bonds rises and alsoto the increase in the viscosity of the starting materials and thereaction mixture. To achieve satisfactory conversions, it is thereforenecessary to employ long residence times, but these adversely affect theselectivity as a result of undesirable secondary reactions, inparticular the formation of high-boiling reaction by-products.Furthermore, the purity demands made of the direct process product arevery high because, due to its high viscosity, it is virtually impossibleto fractionate and thus has to be used essentially without furtherpurification. A reduction in the reaction volume is of particularimportance for the economics of the process, since large reactionvolumes are very expensive at the high pressures needed in the process.

To keep the reaction volume as small as possible and to increase thespace-time yield, it has therefore been proposed that a reactor cascadebe used in place of a continuous stirred tank reactor (CSTR). Cascadingcan be achieved by various methods, but these have the disadvantagesdescribed below:

customary method is cascading by connecting reactors in series. However,this method is the most expensive since the capital cost of theindividual reactors, the associated connections and the instrumentationrequired increase with the number of reactors.

WO 97/20793 describes a process for carrying out hydroformylations in asingle reactor in which cascading is provided by horizontally arrangedperforated plates. The configuration of the stirring elements of thecentrally arranged stirrer are matched to the openings and intermediatespaces of the perforated plates. Such a reactor has, in particular, thefollowing disadvantages: high-pressure apparatuses are generallyconstructed as cylinders whose ends have a lower diameter than thecentral part, i.e. as double-necked apparatuses. It is very difficult inengineering terms to install horizontal perforated plates whose diametercorresponds to that of the middle section through such necks. Cascadingby means of horizontal perforated plates has the further disadvantagethat these perforated plates maintain a high pressure in the event of anemergency depressurization and therefore have to be designed with verygreat thicknesses. Further, the impeller shaft extending into thereaction chamber has to be inserted pressure-tight.

It is an object of the present invention to provide a process for thehydroformylation of olefins having at least 6 carbon atoms which gives avery high conversion at a given reactor volume. A further object is toprovide an apparatus for carrying out the process which has low capitaland operating costs.

We have found that this object is achieved by a continuous process forthe hydroformylation of olefins having at least 6 carbon atoms in thepresence of a homogeneous catalyst, wherein

-   a) a vertical tall cylindrical reactor (1) whose interior space is    divided by means of internals (2) into at least two reaction    chambers which extend essentially in the longitudinal direction of    the reactor is used,-   b) at least one olefin is introduced into the reactor together with    synthesis gas at the lower end of the first reaction chamber,-   c) a partially reacted reaction mixture is conveyed from the upper    end of a reaction chamber to the lower end of a next reaction    chamber; and-   d) the hydroformylated olefin is taken off at the upper end of the    last reaction chamber.

Olefins having at least 6, e.g. from 9 to 20, carbon atoms can behydroformylated by the process of the present invention. For practicalreasons, the maximum chain length is generally limited to about 700carbon atoms. The process of the present invention is particularlyuseful for the hydroformylation of isomeric olefin mixtures prepared byoligomerization of lower olefins such as propene and butenes. Typicaloligomers which are suitable as starting materials for the presentprocess include, inter alia, propene dimer, propene trimer and propenetetramer, butene dimer, butene trimer and butene tetramer and alsocooligomers of propenes and butenes. The oligomers of butenes can beobtained industrially by known oligomerization processes, e.g. by theOctol® process of Hüls and the Dimersol® process of IFP. Furthermore,linear long-chain olefins having a terminal double bond which areobtainable, for example, by the SHOP® process or Ziegler processes orlinear long-chain olefins having an internal double bond can also behydroformylated by the process of the present invention.

Further preferred starting olefins are essentially monounsaturatedpolyalkylenes having from 30 to 700 carbon atoms, in particularpolybutene or polyisobutene.

The process of the present invention is homogeneously catalyzed. Ingeneral, this is achieved by introducing a suitable catalyst or catalystprecursor into the reactor together with the olefin and the synthesisgas. There are no significant restrictions in respect of the catalystsor catalyst precursors which can be used. In particular, cobaltcatalysts, preferably cobalt carbonyls or hydridocobalt carbonyl ortheir precursors, in particular cobalt(II) salts such as cobalt(II)formate, cobalt(II) acetate or cobalt(II) ethylhexanoate, are used in amanner known per se.

The catalyst is advantageously introduced as a solution in the startingolefin or an organic solvent which may be used in addition. For thispurpose, it is possible to bring an aqueous cobalt(II) salt solutioninto contact with synthesis gas outside the reactor to form ahydroformylation-active cobalt catalyst and simultaneously orsubsequently bring the aqueous solution containing the cobalt catalystinto contact with the starting olefin and/or the organic solvent so thatthe cobalt catalyst is extracted into the organic phase.

The prior formation of the catalyst is preferably carried out at from 50to 200° C., in particular from 100 to 160° C., under pressures of from100 to 400 bar, in particular from 200 to 300 bar.

Suitable apparatuses are customary apparatuses for gas/liquid reactions,e.g. stirred vessels provided with a sparging stirrer, bubble columns ortrickle-bed columns. The precarbonylation is advantageously carried outin the presence of activated carbon, zeolites or basic ion exchangersloaded with cobalt carbonyl, as described in DE-A 2139630. The cobaltcatalyst is then extracted from the resulting aqueous solutioncontaining cobalt(II) salts and cobalt catalyst into the olefins to behydroformylated and/or the organic solvent which may be used inaddition.

In many cases it is preferable, in view of the reduced processengineeering costs, to carry out the formation of the cobalt catalyst,the extraction of the cobalt catalyst in the organic phase and thehydroformylation of the olefins in one step by introducing the aqueouscobalt(II) salt solution, the olefins, synthesis gas and any organicsolvent used together into the reactor. The starting materials areintroduced into the reaction zone so that good phase mixing occurs and avery high mass transfer area is generated. For introducing the startingmaterial into the reactor, it is possible to use the introductiondevices known to those skilled in the art, e.g. turbulence tubes filledwith packing elements or mixing nozzles for multiphase systems.

To remove the cobalt catalyst after the reaction, the crude reactionproduct is appropriately depressurized to intermediate pressure, ingeneral from 10 to 30 bar, after leaving the reaction zone and is passedto a cobalt removal stage. In the cobalt removal stage, the reactionproduct is freed of cobalt carbonyl complexes by means of air or oxygen,preferably at from 90 to 130° C., in the presence of an aqueous, weaklyacidic cobalt(II) salt solution. The removal of cobalt can, if desired,be carried out in a pressure vessel filled with packing elements, e.g.Raschig rings, in which a very high mass transfer area is generated. Theorganic product phase is separated from the aqueous phase in adownstream phase separation vessel. In the cobalt removal stage, thehydroformylation-active cobalt catalyst is decomposed to form cobalt(II)salts, predominantly cobalt(II) formate. The aqueous cobalt(II) saltsolution is advantageously returned to the reaction zone or the catalystformation stage.

As an alternative, it is possible to use rhodium catalysts which may bemodified by nitrogen- or phosphorus-containing ligands. The rhodiumcatalyst is generally separated from the hydroformylation product bydistillation, in which the rhodium catalyst remains as residue togetherwith high-boiling constituents.

Synthesis gas is an industrially available mixture of carbon monoxideand hydrogen. The composition of the synthesis gas used in the processof the present invention can vary within a wide range. The molar ratioof carbon monoxide to hydrogen is generally from about 10:1 to 1:10, inparticular from 2.5:1 to 1:2.5. A preferred ratio is about 1:1.5.

As organic solvent which may be used in addition, it is possible to useinert hydrocarbons such as paraffin fractions, aromatic hydrocarbonssuch as benzene, toluene or xylene, or an aldehyde and/or alcohol, inparticular the hydroformylation product of the olefin used. High-boilingby-products of the hydroformylation can also be used as solvent. The useof a solvent may be advantageous, for example, for reducing theviscosity in the case of long-chain olefins.

The temperature in the hydroformylation is generally from 100 to 250°C., in particular from 145 to 200° C. The reaction is preferably carriedout at a pressure in the range from 20 to 400 bar, in particular from200 to 300 bar.

The preheated or unpreheated starting olefin, synthesis gas and, ifappropriate, catalyst or catalyst precursor are fed into a vertical,tall cylindrical reactor which has at least two reaction chambers.According to the present invention, the reaction chambers extendessentially in the longitudinal direction of the reaction and are formedby division of the interior space of the reactor by means of internalswhich are arranged essentially in the longitudinal direction of thereactor. “Essentially in the longitudinal direction of the reactor”means that the reaction chambers have their longest dimension in thedirection of the longitudinal axis of the reactor and the length of areaction chamber is more than 50%, preferably more than 60%, of thelength of the reactor. The internals in each case form a reactionchamber enclosed on all sides with an inlet for the reaction mixture atone end and an outlet for the reaction mixture at the opposite end. Thereaction mixture is thus firstly fed in at the bottom of a firstreaction chamber, flows through this from the bottom upward and at theupper end of this chamber is fed via suitable means for therecirculation of fluid to a second reaction chamber at its lower end.From this second reaction chamber, the reaction mixture can bedischarged from the reactor, but it is also possible to configure thereactor and the process so that the reaction mixture is conveyed fromthe upper end of the second reaction chamber via means for therecirculation of fluid to the lower end of a third reaction chamber, ifdesired flows in an analogous manner through further reaction chambersand is finally discharged from the upper part of the last reactionchamber.

The internals are preferably configured as cylinders which are closed atboth ends and have inlet and outlet openings at opposite ends. There arein principle no restrictions in respect of the arrangement of theinternals in the reactor cross section; the internals can be arrangednext to one another, preferably distributed uniformly over the reactorcross section, but are particularly preferably arranged concentricallyto one another and to the outer wall of the reactor.

In a preferred embodiment, the internals form an internal cylinder whichis closed at both ends and has an inlet at the lower end and an outletat the upper end. The reaction mixture is particularly preferably firstfed into the intermediate space between the interior wall of the reactorand internals, conveyed from the upper end of the intermediate space viameans for the recirculation of fluid to a second reaction chamber at itslower end, if desired conveyed from the upper end of this to furtherreaction chambers through which it is passed from the bottom upward andfinally discharged from the reactor.

In this embodiment, the second reaction chamber thus serves as anafter-reaction zone, i.e. a zone for the completion of the olefinconversion achieved in the first reaction chamber, without undesirablebackmixing with the contents of the first reaction chamber.

In one possible variant, unreacted synthesis gas can be taken, e.g.drawn off under suction, from the gas space at the upper end of one ormore reaction chambers with the exception of the last reaction chamber,compressed and fed back into the reactor from below. This measure can beemployed regardless of the number of reaction chambers present. If onlytwo reaction chambers are present, unreacted synthesis gas is, forexample, taken off from the upper region of the intermediate spacebetween the interior wall of the reactor and internals and fed back intothe same reaction chamber in its lower region. In the case of more thantwo reaction chambers, unreacted gas can be taken off from one or moreof the reaction chambers, in each case in its upper part, but not fromthe reaction chamber which is the last through which the reactionmixture flows. The recirculation of synthesis gas increases theturbulence in the first reaction chamber and ensures intimate contact ofthe gaseous and liquid phases.

The compression of the recirculated, unreacted synthesis gas canadvantageously be carried out by means of a jet pump, in which case itis particularly useful to operate the jet pump by means of the feedstream comprising olefin and freshly introduced synthesis gas. The jetpump can be located in the lower region of the reactor, but it is alsopossible to install it underneath (and outside) the reactor.

In a further embodiment, it is possible for the jet pump to be operatedusing not only the feed stream but also partially reacted reactionmixture which is, for example, drawn off from the reaction at the lowerend of the first reaction chamber by means of a circulation pump. Thereaction mixture which has been drawn off is particularly preferablyfirstly passed through a heat exchanger to remove heat before being fedto the jet pump.

Between the individual reaction chambers in the interior of the reactor,the reactor is provided with means for the recirculation of fluid whichin each case convey the reaction mixture from the upper end of areaction chamber to the lower end of a next reaction chamber. Thesemeans for the recirculation of fluid are particularly preferablyconfigured as downpipes or siphons. The downpipes are configured so thattheir upper edge keeps the liquid level, i.e. the position of theliquid/gas phase boundary, in the respective reaction chamber constant.Should the liquid level drop below the upper edge of the downpipe, thegas pressure which builds up causes more unreacted synthesis gas to bepassed through the downpipe into the next reaction chamber until theliquid level once again reaches the upper end of the downpipe, and viceversa. In this way, control of the level is achieved without complicatedengineering measures.

The tall cylindrical reactor preferably has an aspect ratio, i.e. aratio of length to diameter, I/d, of from about 3:1 to 30:1, inparticular from about 5:1 to 10:1.

The internals in the reactor are preferably dimensioned so that theratio of the volume of the first reaction chamber to that of the secondreaction chamber or of the volumes of two further successive reactionchambers is from 4:1 to 1:4, preferably from 1.5:1 to 1:1.5,particularly preferably about 1:1.

The dimensions of the internals in the longitudinal direction aredesigned so that the height of the second reaction chamber (of thefurther reaction chambers) is more than 50%, in particular more than60%, of the height of the reactor, particularly preferably from 80 to95% of the height of the reactor.

In a preferred embodiment, additional cascading of the last reactionchamber, e.g. the second reaction chamber, can be achieved by means ofhorizontal diffusion-inhibiting devices, e.g. perforated plates whichare preferably arranged so as to be equidistant from one another. Thenumber of perforated plates can be from 1 to 10, particularly preferablyfrom 3 to 5. This measure has a conversion-promoting effect on thehydroformylation reaction, so that a further increase in conversion canbe achieved for a given reaction volume. The horizontal perforatedplates can advantageously be mounted in the internals forming thefurther reaction chambers before these internals are installed in thereactor, so that assembly costs are significantly reduced. They can bemounted in a removable fashion by means of appropriate flangeconnections or else can be welded directly onto the internals.

The interior of the reactor of the present invention can be providedwith means for the indirect cooling of the reaction chambers. Such meansare particularly preferably provided in the first reaction chamberthrough which the reaction mixture flows. They are preferably in theform of cooling coils through which a cooling medium flows. The coolingcoils can be fixed to the internals, e.g. welded onto them, or beinstalled at a distance from the internals.

The invention is illustrated below with the aid of a drawing andexamples:

In the drawing:

FIG. 1 schematically shows a longitudinal section of a preferredembodiment of an apparatus according to the present invention,

FIG. 2 shows a longitudinal section of a further preferred embodiment ofan apparatus according to the present invention with recirculation ofgas by means of a jet pump,

FIG. 3 shows a further embodiment of the apparatus from FIG. 2 withadditional recirculation of part of the reaction mixture via the jetpump,

FIG. 4 shows a longitudinal section of a further preferred embodimentwith cooling coils fixed to the internals,

FIG. 5 shows a variant with cooling coils which are not fixed to theinternals,

FIG. 6 shows both a longitudinal section and a cross section (FIG. 6 a)of a variant with three reaction chambers and

FIG. 7 a shows a cross section through a reactor according to thepresent invention having no concentric internals.

In the figures, identical reference numerals are used to denote the samefeatures.

FIG. 1 shows, by way of example, a longitudinal section of a reactor 1according to the present invention having a neck 8 at each end of thereactor, internals 2 which form a concentric internal cylinder closed atboth ends and an inlet 4 into the intermediate space 3 between theinterior wall of the reactor and the internals 2 and an outlet 5 fromthe reaction chamber bounded by the internals 2, and having means 6 forthe recirculation of fluid from the upper part of the intermediate spacebetween the interior wall of the reactor and internals 2 to the lowerend of the reaction chamber bounded by the internals 2 and alsohorizontal perforated plates 7, with seven perforated plates beingdepicted by way of example.

FIG. 2 shows a longitudinal section of a further preferred embodimentwhich additionally has a return line for gas 9 from the upper region ofthe intermediate space between the interior space of the reactor andinternals 2, with the return line 9 leading to a jet pump 10 which isinstalled underneath the reactor and is operated by means of the liquidfeed mixture.

FIG. 3 shows a longitudinal section of a further preferred embodiment inwhich the jet pump 10 located underneath the reactor is additionallydriven by liquid reaction mixture which is taken off from the lower partof the reactor by means of a circulation pump 11 and fed to the jetpump, and the liquid reaction mixture is firstly cooled by means of aheat exchanger 12.

The embodiment shown in longitudinal section in FIG. 4 additionally hascooling coils 13 which are fixed to the internals 2 and thus cool boththe intermediate space between the interior wall of the reactor andinternals 2 and the interior space of the reactor bounded by theinternals 2.

In contrast, the variant shown in FIG. 5 has cooling coils 13 which arelocated only in the intermediate space between the interior wall of thereactor and internals 2 and thus cool predominantly this intermediatespace.

FIG. 6 shows an illustrated embodiment having three reaction chamberswhich are formed by concentric arrangement of two cylindrical internals2 within the reactor 1. The cross section A/A through the reactor ofFIG. 6 (FIG. 6 a) clearly shows the concentric arrangement of theinternals 2.

On the other hand, FIG. 7 a shows a cross section through a reactorhaving three internals 2 which are not installed concentrically in thereactor 1.

The present invention thus provides a process which gives an improvedspace-time yield and achieves an increase in conversion and selectivity.The apparatus of the present invention has lower capital and operatingcosts than do known apparatuses.

A particularly advantageous aspect is the improvement in the masstransfer between liquid reaction medium and gas phase as a result of thedivision according to the present invention of the reactor cross sectioninto a plurality of zones by means of the internals arranged essentiallyin the longitudinal direction of the reactor. This achieves an increasein the cross-sectional throughput in the zones, resulting in improvedmass transfer. In this way, the very low gas content and mass transferareas in known apparatuses can be increased by a factor of two or morewithout the use of rotating parts, for example stirrers. Thedisadvantages associated with the use of rotating parts, particularly inthe case of high-pressure apparatuses, can thus be avoided.

EXAMPLE

The invention is illustrated by the following mathematical simulation.The simulation is based on a kinetic model of the hydroformylation ofpolyisobutene, which has been set up by mathematical fitting of a largenumber of experimental measurements. The simulation was carried out fora mixture of 80% by weight of a reactive, only slightly branchedpolyisobutene A₁ and 20% by weight of a less reactive, more stronglybranched polyisobutene A₂ as feed. A 10% excess of synthesis gas (molarratio of CO/H₂=1:1), a reaction temperature of 130° C. and a pressure of280 bar were employed as basis for the simulation. It was assumed thatthe reaction rates of the reaction of A₁ and A₂ to form the respectivehydroformylation product are proportional to the mole fractions of A₁and A₂, respectively. Any dependence on the CO or H₂ concentration wasdisregarded. The simulation was carried out on the basis of the reactionvolume remaining unchanged in the reaction. The mass flow of olefin feedand catalyst solution (aqueous cobalt(II) formate solution) was in eachcase set to 6000 kg/h.

In case I, the simulation was carried out for a reactor withoutinternals and having a volume of 30 m³ and total backmixing. In case II,the reactor was divided into four ideally mixed substages having volumesof 15 m³, 5 m³, 5 m³ and 5 m³. The inlet concentrations of the olefinfeed and the amounts of catalyst are the same in both reactorconfigurations.

The results obtained are summarized in the following table.

Conversion Conversion Total conversion Example of A₁ of A₂ of A₁ + A₂ I78.3% 26.5% 67.9% II 91.3% 28.9% 78.8%

It can be seen that both the individual conversions and the totalconversion are higher in the case of the cascading according to thepresent invention of the reactor at the same reaction volume.

1. A continuous process for the hydroformylation of olefins having at least 6 carbon atoms in the presence of a homogeneous catalyst, wherein a) a vertical tail cylindrical reactor whose interior space is divided by means of internals into at least two reaction chambers which extend essentially in the longitudinal direction of the reactor is used, b) at least one olefin is introduced into the reactor together with synthesis gas at the lower end of the first reaction chamber, c) a partially reacted reaction mixture is conveyed from the upper end of a reaction chamber to the lower end of a next reaction chamber; and d) the hydroformylated olefin is taken off at the upper end of the last reaction chamber.
 2. A process as claimed in claim 1, wherein a reactor whose interior space is divided by means of internals so that the second and any further reaction chambers are arranged essentially concentrically to the outer wall of the reactor is used.
 3. A process as claimed in claim 1, wherein unreacted synthesis gas is taken from the gas space at the upper end of one or more reaction chambers with the exception of the last reaction chamber and is fed back into the reactor.
 4. A process as claimed in claim 3, wherein the synthesis gas which has been taken off is fed back into the reactor by means of a jet pump which is operated by means of the olefin and synthesis gas fed in.
 5. A process as claimed in claim 4, wherein the jet pump is additionally operated by means of a partially reacted reaction mixture which is taken off at the lower end of the first reaction chamber. 